Sulfur oxide and particulate removal system

ABSTRACT

A system is provided in which particulates and sulfur oxides are simultaneously removed from flue gases in a granular bed filter with special sulfur oxide-capturing and particulate-removing material. The spent sulfur oxide-capturing and particulate-removing material can be regenerated in a lift pipe riser.

BACKGROUND OF THE INVENTION

This invention relates to flue gas cleanup and, more particularly, toremoving sulfur oxides and particulates from a gaseous stream, such asfrom a regenerator in a catalytic cracking unit.

Flue gases emitted in combustors, such as in regenerators and powerplants, often contain undesirable levels of sulfur oxides (SOx),nitrogen oxides (NOx), and particulates which, if untreated, mightpollute the atmosphere.

Sulfur oxides in the presence of water can form sulfuric acid causingacid rain. Nitrogen oxides may cause smog by photochemical reaction withhydrocarbons in the atmosphere. Particulates in flue gases typicallyinclude ash (soot) and/or spent combusted catalyst with trace metals,such as arsenic and other contaminants which, in excessive levels, couldpoison vegetation and livestock.

Over the years, various methods have been suggested for controllingand/or removing sulfur oxide and/or nitrogen oxide emissions. Incatalytic cracking units, sulfur oxide control processes usually occurin the regenerator. In one widely used process, sulfur oxides arereacted in the regenerator in the presence of a catalyst to formhydrogen sulfide which is withdrawn with the product stream from thecatalytic cracker and subsequently treated in a sulfur recovery plant.Some of the methods suggested for removing nitrogen oxides inregenerators, however, poison the cracking catalyst and are, therefore,unacceptable. Typifying these prior art methods for controlling sulfuroxide and/or nitrogen oxide emissions are those described in U.S. Pat.Nos. 2,493,218; 2,493,911; 2,522,426; 2,575,520; 2,863,824; 2,992,895;3,023,836; 3,068,627; 3,264,801; 3,501,897; 3,755,535; 3,760,565;3,778,501; 3,832,445; 3,835,031; 3,840,643; 3,846,536; 3,892,677;4,001,376; 4,006,066; 4,039,478; 4,153,534; 4,153,535; 4,181,705;4,206,039; 4,218,344; 4,221,677; 4,233,276; 4,238,317; 4,241,033;4,254,616; 4,258,020; 4,267,072; 4,300,997; 4,323,542; 4,325,811;4,369,109; 4,369,130; 4,376,103; 4,381,991; 4,405,443; 4,423,019; and4,443,419. These prior art methods have met with varying degrees ofsuccess.

Flue gas streams discharged from regenerators, power plants, or othercombustors are commonly directed through one or more dedusters, such asflue gas scrubbers, electrostatic precipitators, cyclones, bag houses,granular bed filters, or other filters, in order to remove particulatesfrom the flue gas stream. Typifying these dedusters and other prior artparticulate-removing devices are those shown in U.S. Pat. Nos.3,540,388; 3,550,791; 3,596,614; 3,608,529; 3,608,660; 3,654,705;3,672,341; 3,696,795; 3,741,890; 3,769,922; 3,818,846; 3,882,798;3,892,658; 3,921,544; 3,922,975; 4,017,278; 4,126,435; 4,196,676; and4,421,038. These dedusters and prior art devices have met with varyingdegrees of success.

The combined use of flue gas scrubbers and electrostatic precipitators,while often effective to control particulate emissions, is veryexpensive and cumbersome.

It is therefore desirable to provide an improved system to remove sulfuroxides and particulates from gaseous streams.

SUMMARY OF THE INVENTION

An improved system is provided for efficiently, effectively, andeconomically removing sulfur oxides (SOx) and particulates from gaseousstreams, such as flue gases, to minimize emission of pollution andcontaminants into the atmosphere. The novel system is particularlyuseful to clean up combustion off-gases emitted from regenerators ofcatalytic cracking units to environmentally acceptable levels. Thesystem is also beneficial to effectively remove sulfur oxides andparticulates from combustion gases emitted from synthetic fuel plants,such as those which retort, solvent extract, or otherwise process oilshale, tar sands, diatomaceous earth (diatomite), uintaite (gilsonite),lignite, peat, and biomass, as well as to effectively remove sulfuroxides and particulates emitted from coal liquefaction and gasificationplants. The disclosed system is also useful to clean up flue gases frompower plants, paper mills, steel mills, waste (garbage) treatment sites,chimneys, smoke stacks, etc. The system is also useful for removingnitrogen oxides (NOx) from gaseous streams.

To this end, sulfur oxide, nitrogen oxide, and particulate-laden gasesare treated and purified in a single processing vessel, preferably agranular bed filter, located downstream of the combustor to removesimultaneously sulfur oxides, nitrogen oxides, and particulates from thegases. In the processing vessel, the particulates, nitrogen oxides, andsulfur oxides are simultaneously removed from the dusty sulfur andnitrogen oxide-containing gases by passing the gases through at least aportion of a bed of sulfur oxide-capturing, nitrogen oxide-capturing,and particulate-removing material. Desirably, the gases are fed into thevessel and passed through the portion of the bed at an angle ofinclination from 30° to 90° relative to the horizontal axis of thevessel and most preferably vertically downwardly at right angles(perpendicular) to the horizontal axis for best results.

Preferably, the bed of sulfur oxide-capturing, nitrogen oxide-capturing,and particulate-removing material is a downwardly moving bed of granularmaterial in the form of balls, spheres, pebbles, or pellets. Thepreferred granular material is alumina adsorbers, although adsorberscomprising oxides of other metals can also be used, either alone or incombination with alumina and/or each other, such as bismuth, manganese,yttrium, antimony, tin, rare earth metals, Group 1a metals, and/or Group2a metals.

The adsorbers can be coated with a catalyst that promotes the removal ofsulfur oxides. While the preferred catalyst is platinum, other catalyticmetals, both free and in a combined form, can be used, either alone orin combination with platinum and/or each other, such as rare earthmetals, Group 8 metals, chromium, vanadium, rhenium, and combinationsthereof.

The spent material (adsorbers) containing the captured particulates andsulfur oxides can be regenerated, such as in a lift pipe riser ortransfer line, to remove the sulfur oxides and particulates from theadsorbers. The regenerated adsorbers can be recycled to the processingvessel, with or without additional scrubbing or stripping, as desired.In one form, the adsorbers are regenerated thermally, such as bycombustion or other heating means. In another form, the adsorbers areregenerated with a reducing gas to convert the sulfur oxides to hydrogensulfide. The reducing gas can be hydrogen, ammonia, carbon monoxide, orlight hydrocarbon gases, such as methane, ethane, propane, etc. Thereducing gas can also be diluted with steam to attain a shift reactionor steam reforming in order to produce hydrogen and carbon dioxide.Hydrogen produced in this manner serves as a very effective andrelatively inexpensive reducing gas to convert sulfur oxides to hydrogensulfides. The hydrogen sulfide can be treated in a hydrogen sulfidetreatment plant, such as an amine recovery unit and a Claus plant. Theparticulates in the dusty effluent gases can be removed downstream ofthe regenerator in one or more filters, such as a cyclone and/or baghouse.

Particulates emitted from catalytic cracking units are mainly combustedspent catalyst. Particulates emitted from synthetic fuel plants aremainly combusted synthetic fuels (spent hydrocarbon-containingmaterial). Particulates emitted from power plants, steel mills, wastetreatment sites, etc., contain ash and/or other material.

As used in this application, the terms "sulfur oxide" and "sulfuroxides" mean sulfur dioxide and/or sulfur trioxide.

The term "SOx" as used herein means sulfur oxide.

The terms "nitrogen oxide" and "nitrogen oxides" as used herein meannitric oxide (NO) and/or nitrogen dioxide (NO₂).

The term "NOx" as used herein means nitrogen oxide.

The terms "spent catalyst,""spent promoter," and "spent material" asused herein mean a catalyst, promoter, or material, respectively, whichhas been at least partially deactivated.

A more detailed explanation of the invention is provided in thefollowing description and appended claims taken in conjunction with theaccompanying drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic flow diagram of a gas purification system inaccordance with principles of the present invention;

FIG. 2 is a schematic flow diagram of part of the gas purificationsystem with air, instead of a reduction gas, being injected into thelift pipe riser;

FIG. 3 is a schematic flow diagram of an amine recovery unit;

FIG. 4 is a schematic flow diagram of a sulfur recovery unit; and

FIG. 5 is a cross-sectional view of a catalytic cracking unit.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT

Referring now to FIG. 1, a gas purification system 10 is provided toremove sulfur oxides (SOx) and particulates from a gaseous stream 12,such as flue gases, to minimize emission of pollution and contaminantsinto the atmosphere. While the system of the present invention isdescribed hereinafter with particular reference to cleanup of combustionoff-gases emitted from the regenerator or combustor 14 of a catalyticcracking unit 16, it will be apparent that the system of the presentinvention can also be used to effectively clean up combustion gases(flue gases) emitted from other combustors, such as those from syntheticfuel plants, which retort, solvent extract, or otherwise process oilshale, tar sands, diatomaceous earth (diatomite), uintaite (gilsonite),lignite, peat, and biomass, coal liquefaction and gasification plants,power plants, paper mills, steel mills, waste (garbage) treatment sites,chimneys, smoke stacks, etc.

In the gas purification system of FIG. 1, a hydrocarbon feedstock, suchas gas oil, is fed through feedstock line 18 into the bottom of acatalytic cracking reactor 20, such as a fluid catalytic cracker (FCC).Fresh make-up catalytic cracking catalyst and regenerated catalyticcracking catalyst are fed into the reactor through fresh make-upcatalyst line 22 and regenerated catalyst line 24, respectively. In thereactor, the hydrocarbon feedstock is volatilized, gasified, andfluidized upon being mixed with the hot cracking catalyst and thefeedstock is catalytically cracked to more valuable hydrocarbons. Thecatalytically cracked hydrocarbons are withdrawn from the top of thereactor through overhead product line 26 and conveyed to downstreamprocessing equipment (not shown) for further upgrading, separation intofractions, and/or further processing.

Spent catalyst is discharged from the reactor through spent catalystline 28 and fed to the bottom portion of an upright, fluidizablecatalyst regenerator or combustor 14. The reactor and regeneratortogether provide the primary components of the catalytic cracking unit.Air is injected upwardly into the bottom portion of the regeneratorthrough air injector line 30 by air pump 32. The air is injected at apressure and flow rate to fluidize, propel, convey, and transport thespent catalyst particles generally upwardly through the regenerator.Residual carbon (coke) contained on the catalyst particles issubstantially completely combusted in the regenerator leavingregenerated catalyst for use in the reactor. The regenerated catalyst isdischarged from the regenerator through regenerated catalyst line 24 andfed to the reactor. The combustion off-gases (flue gases) are withdrawnfrom the top of the combustor through an overhead combustion off-gasline or flue gas line 12. The combustion off-gases or flue gases containminute particulates of combusted spent catalyst particles as well assulfur oxides (SOx) and nitrogen oxides (NOx). The particulates in thecombustion off-gases and flue gases emitted from the regenerator of acatalytic cracking unit are very small and typically range in size from20 microns to less than 0.1 micron. Under present governmentenvironmental standards, the particulates, SOx, and NOx in the fluegases are pollutants which must be reduced to environmentally acceptablelevels before the flue gases are vented to the atmosphere.

In the regenerator 33 of FIG. 5, the regenerator has a dense phase lowersection 34 and a dilute phase upper section 36 to provide forsubstantially complete carbon monoxide (CO) burning and combustion inthe manner described by Horecky et al., U.S. Pat. No. 3,909,392, whichis hereby incorporated by reference in its entirety. The regenerator canalso have one or more internal cyclones 38 and 39 for removing some ofthe combusted particulates from combustion gases. The removed catalystparticles are discharged through dip legs or return lines 40 and 41 atthe lower end of the cyclones into the dense phase lower portion 34. Theregenerator can also have an eductor or eductor tube 42 to disperse thespent cracking catalyst particles in a fountain, rain, or spouted bedinto the dilute phase upper portion of the regenerator, via valve 43,with the aid of air, steam, or inert gases.

As shown in FIG. 5, the catalytic cracking reactor 43 (catalyticcracker) can also have a dense phase lower portion 44 and a dilute phaseupper portion 46, as well as one or more internal cyclones 48 and 49 forremoving cracking catalyst particles from the gaseous product streambefore the cracked product stream is removed from the reactor.Downwardly depending dip legs or return lines 50 and 51 from theinternal cyclones in the reactor return the cracking catalyst particlesto the lower portion of the reactor. If desired, external cyclones canbe used instead of internal cyclones.

The reactor 43 can also have a steam stripping section 52 at the bottomof the reactor which is of a smaller cross-sectional area than the outerwalls of the dilute and dense phase portions of the reactor. Steam isinjected into the steam stripping portion 52 through steam line 54 tosteam strip volatile hydrocarbons from the cracking catalyst particles.The steam also serves to fluidize the cracking catalyst in the strippingportion 52 as well as to fluidize the cracking catalyst in the lowerdense phase 44 of the reactor. The steam stripping portion can haveinternals, such as conical baffles 56 and donuts 57, to enhance flow andsteam stripping. A high temperature second stage steam stripper can alsobe used.

The spent catalyst can be withdrawn from the bottom of the steamstripper section through spent catalyst line 58, via control valve 59,instead of from the upper portion of the reactor, if desired, and can betransported upwardly into the lower portion of the regenerator 33through a transfer line 60 and regenerator inlet lines 61 and 62, viainlet valves 63 and 64, with the aid of air from air injector 65. Theregenerated catalyst can be withdrawn from the bottom of the regenerator33 through regenerated catalyst lines 66a and 66b, if desired, insteadof from the upper portion of the regenerator and conveyed by regeneratedcatalyst line 67, valve 68, and reactor inlet line 69 to the dilutephase portion 46 of the reactor 43 along with the hydrocarbon feedstockfrom feedstock line 70. The temperature in the regenerator can becontrolled by steam pod injector 71.

Suitable hydrocarbon feedstocks for the catalytic cracking unitpreferably have a boiling point above the gasoline boiling range, forexample from about 400° F. to about 1,200° F., and are usuallycatalytically cracked at temperatures ranging from about 850° F. toabout 1,200° F. Such feedstocks can include various mineral oilfractions boiling above the gasoline range, such as light gas oils,heavy gas oils, wide-cut gas oils, vacuum gas oils, kerosenes, decantedoils, residual fractions (resid), reduced crude oils, and cycle oilsderived from any of these, as well as suitable fractions derived fromshale oil, tar sands oil, diatomaceous earth oil, coal liquefaction, orother synthetic oils. Such fractions may be employed singly or in anydesired combination.

Catalytic cracking of heavy mineral oil fractions is an importantrefining operation in the conversion of crude oils to desirable fuelproducts, such as high-octane gasoline fuel used in spark-ignited,internal combustion engines. In fluid catalytic cracking, high molecularweight hydrocarbon liquids or vapors are contacted with hot,finely-divided solid catalyst particles in a fluidized bed reactor suchas shown in FIG. 1 or in an elongated riser reactor such as shown inFIG. 5, and the catalyst-hydrocarbon mixtures are maintained at anelevated temperature in a fluidized or dispersed state for a sufficientperiod of time to obtain the desired degree of cracking to lowermolecular weight hydrocarbons typically present in motor gasoline anddistillate fuels.

In usual cases where riser cracking is employed for conversion of a gasoil, the throughput ratio, or volume of total feed to fresh feed, mayvary from about 1 to 3. The conversion level may vary from about 40 toabout 100 weight percent and advantageously is maintained above about 60weight percent, for example, between about 60 and 90 weight percent. Theterm "conversion" is generally used herein as the percentage reductionby weight of hydrocarbons boiling above about 430° F. at atmosphericpressure by the formation of lighter materials or coke.

The weight ratio of total cracking catalyst-to-oil in the riser reactor(catalytic cracker) can vary within the range of from about 2 to about20 in order that the fluidized dispersion will have a density within therange of from about 1 to about 20 pounds per cubic foot. Desirably, thecatalyst-to-oil ratio is maintained within the range of from about 3 toabout 20, preferably 3 to about 7 for best results. The fluidizingvelocity in the riser reactor (catalytic cracker) can range from about10 to about 100 feet per second. The riser reactor can have a ratio oflength-to-average diameter of about 25.

For production of a typical naphtha product, the bottom section mixingtemperature within the riser reactor (catalytic cracker) above thestripping section is advantageously maintained at about 1,000° F. toabout 1,100° F. for vaporization of the oil feed so that the topsection's product exit temperature will be about 950° F. For crackingresids and synthetic fuels, substantially higher mixing temperatures inthe bottom section of the reactor, such as about 2,000° F., may benecessary for effective cracking.

Under the above conditions, including provision for a rapid separationof spent catalyst from effluent oil vapor, a very short period ofcontact between the cracking catalyst and oil will be established.Contact time within the riser reactor (catalytic cracker) will generallybe within the range of from about 1 to about 15 seconds, preferablywithin the range of from about 3 to about 10 seconds. Short contacttimes are preferred because most of the hydrocarbon cracking occursduring the initial increment of contact time and undesirable secondaryreactions are avoided. This is especially important if higher productyield and selectivity, including lesser coke production, are to berealized.

Short contact time between cracking catalyst particles and oil vaporscan be achieved by various means. For example, cracking catalyst may beinjected at one or more points along the length of a lower, or bottom,section of the riser reactor (catalytic cracker). Similarly, oil feedmay be injected at all the points along the length of the lower sectionof the riser reactor and a different injection point may be employed forfresh and recycle feed streams. Auxiliary nozzles can also be used todisperse resids or other feedstock onto the catalyst for more efficientcatalytic cracking reactions. The lower section of the riser reactorabove the stripping section may, for this purpose, include up to about80 percent of the total riser length in order to provide extremely shorteffective contact times inducive to optimum conversion of petroleumfeeds. The reactor is preferably designed to minimize cracking of theproduct in the dilute phase. Where a dense catalyst bed is employed,provision may also be made for injection of cracking catalyst particlesand/or oil feed directly into the dense-bed zone.

While the conversion conditions specified above are directed to theproduction of gasoline as fuel for spark-ignition internal combustionengines, the process may be suitably varied to permit maximum productionof heavier hydrocarbon products such as jet fuel, diesel fuel, heatingoil and chemicals and, in particular, olefins and aromatics.

In catalytic cracking, some non-volatile carbonaceous material, or"coke", is deposited on the catalyst particles. Coke comprises highlycondensed aromatic hydrocarbons which generally contain a minor amountof hydrogen, such as from about 4 to about 10 weight percent. When thehydrocarbon feedstock contains organic sulfur compounds, the coke alsocontains sulfur and nitrogen. As coke builds up on the catalyst, theactivity of the catalyst for cracking and the selectivity of thecatalyst for producing gasoline blending stocks diminish. The catalystparticles may recover a major proportion of their original capabilitiesby removal of most of the coke therefrom in the catalyst regenerator.

The spent catalyst from the petroleum conversion reaction in the reactoris preferably stripped in the steam stripping section 52 (FIG. 5) priorto entering the regenerator. The stripping section for use in thefluidized bed catalytic cracker may be maintained essentially at aconversion reactor temperature in the range of from about 200° F. toabout 1,200° F. and preferably above about 870° F. for best results. Thepreferred stripping gas is steam although steam containing a diluent,such as nitrogen or some other inert gas or flue gas, may also beemployed. The stripping gas can be injected into the stripping sectionat a pressure of at least about 10 psig, preferably about 35 psig, toattain substantially complete removal of volatile compounds from thespent conversion catalyst. If desired, an inert stripping gas may beused instead of steam.

Catalyst regeneration is accomplished by burning the coke deposits fromthe catalyst surface with a molecular oxygen-containing gas, such asair. The oxidation of coke may be characterized in a simplified manneras the oxidation of carbon as shown below.

    C+O.sub.2 →CO.sub.2                                 (a)

    2C+O.sub.2 →2CO                                     (b)

    2CO+O.sub.2 →2CO.sub.2                              (c)

Reactions (a) and (b) both can occur under typical catalyst regenerationconditions with the catalyst temperature ranging from about 1,050° F. toabout 1,450° F. or higher and are exemplary of gas-solid chemicalinteractions when regenerating cracking catalyst at temperatures withinthis range. The effect of any increase in temperature is reflected in anincreased rate of combustion of carbon and a more complete removal ofcarbon, or coke, from the catalyst particles. As the increased rate ofcombustion is accompanied by an increased evolution of heat wheneversufficient free or molecular oxygen is present, the gas-phase reaction(c) may occur. This latter reaction is initiated and propagated by freeradicals and can be catalyzed.

The burning of sulfur-containing and nitrogen-containing coke depositsfrom the catalyst also results in the formation of sulfur oxides andnitrogen oxides, and although the disclosed invention is not to belimited thereby, sulfur-compound and sulfur oxide burning may berepresented by the following chemical equations:

    S (in coke)+O.sub.2 →SO.sub.2                       (d)

    2SO.sub.2 +O.sub.2 →2SO.sub.3                       (e)

Reactions (d) and (e) also occur under typical cracking catalystregeneration conditions. While reaction (d) is fast, reaction (e) isrelatively slow. Reaction (e) can be catalyzed by any catalyst whichcatalyzes reaction (c) above.

Stripped deactivated cracking catalyst is regenerated by burning thecoke deposits from the catalyst surface with air or some othercombustion-sustaining molecular oxygen-containing regeneration gas in aregenerator. This burning results in the formation of combustionproducts such as sulfur oxides, carbon monoxide, carbon dioxide, andsteam. The oxygen-containing regeneration gas can contain a diluent,such as nitrogen, steam, carbon dioxide, recycled regenerator effluentgases, and the like. The molecular oxygen concentration of theregeneration gas is ordinarily from about 2 to about 30 volume percentand preferably from about 5 to about 25 volume percent. Since air isconveniently employed as a source of molecular oxygen, a major portionof the inert gas can be nitrogen. The regeneration zone temperatures areordinarily in the range from about 1,049° F. to about 1,454° F. and arepreferably in the range from about 1,148° F. to about 1,355° F. Otherregeneration temperatures may be used in some circumstances. When air isused as the regeneration gas, it can be injected into the bottom of theregenerator from a blower or compressor at a fluidizing velocity in therange from about 0.15 to about 5 feet per second and preferably fromabout 0.5 to about 3 feet per second.

Suitable cracking catalysts include those containing silica and/oralumina, including the acidic type. The cracking catalyst may containother refractory metal oxides such as magnesia or zirconia. Preferredcracking catalysts are those containing crystalline aluminosilicates,zeolites, or molecular sieves, in an amount sufficient to materiallyincrease the cracking activity of the catalyst, e.g., between about 1and about 25% by weight. The crystalline aluminosilicates can havesilica-to-alumina mole ratios of at least about 2:1, such as from about2 to 12:1, preferably about 4 to 6:1 for best results.

The crystalline aluminosilicates are usually available or made in sodiumform and this component can be reduced, for instance, to less than about4 or even less than about 1% by weight through exchange with hydrogenions, hydrogen-precursors such as ammonium ions, or polyvalent metalions. Suitable polyvalent metals include calcium, strontium, barium, andthe rare earths such as cerium, lanthanum, neodymium, andnaturally-occurring rare earths and their mixtures. Such crystallinematerials are able to maintain their pore structure under the hightemperature conditions of catalyst manufacture, hydrocarbon processingand catalyst regeneration. The crystalline aluminosilicates often have auniform pore structure of exceedingly small size with thecross-sectional diameter of the pores being in a size range of about 6to 20 angstroms, preferably about 10 to 15 angstroms. Silica-basedcracking catalysts having a major proportion of silica, e.g., about 60to 90 weight percent silica and about 10 to 40 weight percent alumina,are suitable for admixture with the crystalline aluminosilicate or foruse as such as the cracking catalyst. Other cracking catalysts and poresizes can be used.

The cracking catalyst particles are finely-divided and may have anaverage particle size in the range of about 150 microns to about 20microns or less.

The regeneration gas serving to fluidize the dense bed contains free ormolecular oxygen and the oxygen is preferably charged to the regeneratorin an amount somewhat in excess of that required for complete combustionof coke (carbon and hydrogen) to carbon dioxide and steam. The amount ofoxygen in excess of that required for complete combustion of the cokemay vary from about 0.1 to about 35 or more percent of the theoreticalstoichiometric oxygen requirement for complete combustion of the cokebut, advantageously, need not be greater than about 10 percent. Forexample, when air is employed as the regeneration gas, a 10 percentexcess of air provides only about 2 volume percent oxygen in theeffluent spent gas stream. Advantageously, the concentration ofmolecular or free oxygen and carbon monoxide at any point within theregenerator is maintained outside of the explosive range at thoseconditions to eliminate any risk of detonation.

An oxidation-promoting, carbon monoxide-burning catalyst can be fed tothe regenerator to promote complete burning of carbon monoxide to carbondioxide in the regenerator. The solid oxidation catalyst can be in afinely-divided form, such as powder, separate from the hydrocarboncracking catalyst, or can be supported on another substrate and admixedwith the cracking catalyst. The support for the oxidation catalyst canbe less catalytically active, or even inert, to the oxidation andhydrocarbon conversion reactions. Desirably, the support is porous. Thesupport can have a surface area, including the area of the pores on thesurface, of at least about 10, preferably at least about 50, squaremeters per gram. Illustrative of the supports, which may be essentiallyamorphous, are silica, alumina, silica-alumina, and the like. Solid,platinum group metal or rhenium oxidation catalysts may be used as canother oxidation catalysts that promote the oxidation of carbon monoxidein the presence of molecular oxygen. These oxidation catalysts contain acatalytic metal which promotes the oxidation. The metal may be incombined form, such as an oxide, rather than being in the elementalstate. The oxidation catalysts can be rhenium or a platinum group metalof Group 8, such as platinum, palladium and rhodium. The solid oxidationcatalyst may comprise two or more catalytically-active metals eitherphysically or chemically combined. By a chemical combination of metals,there are included bi- or poly-metallic salts or oxides. Illustrative ofcombinations of catalytically-active metals which may promote oxidationof carbon monoxide without unduly adversely affecting the hydrocarboncracking operations are combinations of the platinum group metals, e.g.,platinum, rhenium, the oxides of iron and rhenium, and the like. Othermetals can be used.

The substrate for the solid oxidation-promoting carbon monoxide-burningcatalyst may be a portion of the cracking catalyst or may be differenttherefrom, for example, it may be a non-catalytic, porous, solidsubstrate. When the hydrocarbon cracking catalyst serves as thesubstrate, care should be taken in selection of the deposition processsuch that the cracking activity and selectivity of the catalyst is notadversely affected. It is preferred that if the hydrocarbon crackingcatalyst is of the type having ion-exchanged sites, the ion-exchange becompleted prior to deposition of the oxidation catalyst. The amount ofoxidation-promoting metal employed for promotion of the oxidation ofcarbon monoxide may be in a minor amount effective to enhance thedesired oxidation. This amount may be very small, e.g., as little asabout 0.01 part per million or less based on the weight of thehydrocarbon cracking catalyst employed. The amount ofoxidation-promoting metal may often be at least about 0.1 ppm up toabout 5 or about 10 ppm. Larger amounts of the oxidation-promotingmetal, such as about 0.01 to 5, or about 0.05 to 1, percent by weightbased on the hydrocarbon cracking catalyst, may be employed.

In order to remove sulfur oxides and particulates from theregenerator/combustion off-gases (flue gases), the sulfur oxide andparticulate-laden gases are passed through a granular bed filter 80(FIG. 1), either directly or indirectly, after passing through one ormore cyclones to remove some of the large gross particulates. Thegranular bed filter is an elongated, upright single, sulfuroxide-capturing and particulate-removing vessel which filtersparticulates and sulfur oxides from the influent flue gases. Thegranular bed filter has an exterior vertical sidewall 82 with a circularcross-section, an elongated frustro-conical bottom section or portion 84whose flared sidewalls converge downwardly and terminate into an outletor discharge mouth 86 along the vertical axis of the granular bedfilter, and an upwardly converging roof or top 88.

The granular bed filter has an inlet conduit at an angle of inclinationof 30° to about 90° relative to the horizontal axis of the housing ofthe vessel. Preferably, the inlet conduit comprises a vertical conduitor pipe 90 which provides a gas inlet line. The conduit extendsvertically downwardly through the roof along the vertical axis of thefilter to a discharge position in the upper portion of the interior ofthe frustro-conical bottom section. The vertical conduit has an enlargedhead 92 at its upstream end which extends upwardly through the roof, anelongated main body 94 which has a smaller cross-section than the headand is circumferentially surrounded by the sidewall 82, and has anoutwardly flared discharge portion 96 at its downstream end withdownwardly converging frustro-conical walls which terminate in a gasoutlet and discharge mouth 98. The upstream head of the conduitpreferably has a vertical inlet mouth 100, although in somecircumstances it may be desirable to have a horizontal inlet mouth.

Extending downwardly from the roof within the interior of the filter isan annular frustro-conical, adsorber collection reservoir or ball hopper102 whose flared sidewalls converge downwardly and surround the upperportion of the vertical conduit. Discharge chutes or outlet pipes 104and 106 extend generally downwardly from the reservoir into or slightlyabove a downwardly-moving bed 108 of sulfur oxide-capturing andparticulate-removing material. The chutes can include acentrally-disposed vertical discharge chute 104 that circumferentiallyand concentrically surrounds a portion of the main body section 94 ofthe vertical conduit and symmetrical, outwardly inclined, angular chutes106 which extend downwardly and outwardly at an angle of inclinationrelative to the vertical axis of the filter. Extending outwardly fromthe sides of the filter and scrubber is a gas outlet line 110. Aregenerated adsorber-inlet line 114 extends generally downwardly at anangle of inclination through the roof of the filter.

In the preferred embodiment, the bottom frustro-conical section 84 ofthe granular bed filter is filled with a downwardly moving bed 108 ofsulfur oxide-capturing and particulate-removing granular material whichis in the form of balls, pebbles, spheres, or pellets. The sulfuroxide-capturing and particulate-removing material provides adsorbers oracceptors which adsorb, collect, and/or otherwise remove sulfur oxidesand particulates from the influent gaseous stream (regenerator fluegases). In the most preferred embodiment, the bed of granular materialis a bed of sulfur oxide and nitrogen oxide-capturing andparticulate-removing material, which serve as sulfur dioxide, nitrogenoxide, and particulate adsorbers or acceptors. The adsorbers enter thegranular bed filter and scrubber through fresh make-up adsorber line 112or regenerated adsorber line 114 and descend by gravity into thefrustro-conical adsorber reservoir 102. The adsorbers are dischargeddownwardly from the reservoir through the downwardly extending chutesinto the downwardly moving bed.

The adsorbers preferably comprise substantially alumina, and mostpreferably alumina compounded with magnesia, for best results. Gamma (γ)alumina, chi-eta-rho (χ, η, ρ) alumina, delta (δ) alumina, and theta (θ)alumina are particularly useful as adsorbers and supports because oftheir high surface areas. While alpha (α) alumina and beta (β) aluminacan be used as adsorbers, they are not as effective as gamma,chi-eta-rho, delta, and theta alumina. The oxides of other metals canalso be used as adsorbers, either alone or in combination with alumina,or as spinels, such as bismuth, manganese, yttrium, antimony, Group 1ametals, Group 2a metals, rare earth metals, and combinations thereof.Magnesium aluminate spinels are particularly useful as adsorbers.Lanthanum and cerium are preferred rare earth metals. Naturallyoccurring rare earths, such as in the form of baestanite, are alsouseful adsorbers. The adsorbers can also be a blend/mixture of highdensity and low density materials, such as of the above-identified metaloxides.

The adsorbers can be impregnated or otherwise coated with a catalyst orpromoter that promotes the removal of sulfur oxides. The preferredcatalyst is ceria (cerium oxide) and most preferably platinum for bestresults. Other catalytic metals, both free and in a combined form, canbe used, either alone or in combination with each other and/or incombination with ceria and/or alumina, such as rare earth metals, Group8 metals, chromium, vanadium, rhenium, and combinations thereof. Thepromoter can comprise the same material as the adsorber. An even uniformdistribution of the promoter is preferred for best results and tominimize adsorber erosion.

The Group 1a metals, Group 2a metals, and Group 8 metals referred to arethose listed in the Periodic Table of the Elements in the Handbook ofChemistry and Physics (54th Edition). Useful Group 1a metals include:lithium, sodium, potassium, rubidium, and cesium. Useful Group 2a metalsinclude magnesium, calcium, strontium, and barium. Useful Group 8 metalsare the Group 8 noble metals (the palladium family of metals) includingruthenium, rhodium, palladium, osmium, iridium, and platinum. The rareearth metals are also referred to as lanthanides. Useful rare earthmetals include cerium, praseodymium, neodymium, promethium, samarium,europium, gadolinium, terbium, dysprosium, holmium, erbium, thulium,ytterbium, and lutetium.

In operation, the regenerator off-gases (flue gases) in gas line 12 passinto vertical conduit 90 and flow vertically downwardly along and aboutthe vertical axis of the granular bed filter until being discharged fromthe mouth 98 of the conduit into the bed of sulfur oxide-capturing andparticulate-removing material (adsorbers). Depending on the velocity andpressure of the flue gases, the gaseous stream (flue gases) will passdownwardly through a portion of the bed before circulating upwardly. Theadsorbers serve to filter, adsorb, or otherwise remove the particulatesand sulfur oxides (SOx) from the flue gases. The cleansed, purified fluegases are withdrawn from the granular bed filter through the inlet mouthof the gas outlet line 110, located above the bed, where the purifiedgases can be safely vented to the atmosphere or conveyed, expanded, fed,and used to drive and propel the turbine blades of a power recoveryturbine 116 or other equipment. The turbine can be connected to drivethe air blower or pump 32.

In use, the granular bed filter has a highly concentrated collectionzone at the exit (mouth) of the vertical conduit where downwardlyflowing flue gas enters the bed of adsorbers, along with a downstreamcounterflow collection region which substantially assures that cleansed(purified) flue gas always exits upwardly through the downwardly movingbed of adsorbers.

The amount of sulfur dioxide (SO₂) adsorbed on a platinumcatalyst/promoter, such as a 2 ppm platinum catalyst on an aluminaadsorber, depends on the amount of catalyst used (space velocity) aswell as the temperature at which the adsorption is done. The amount ofsulfur dioxide adsorbed, measured as breakthrough time, is greatest ateither low temperatures of about 500° F. or high temperatures of about1,200° F. to about 1,400° F. Sulfur dioxide adsorption will occur atintermediate temperatures ranging from 800° F. to 1,100° F. at anacceptable, but lesser, efficiency. Some sulfur dioxide adsorption mayoccur at a temperature as low as 200° F. and as high as 1,600° F. in thegranular bed filter and scrubber.

The spent adsorbers containing or coated with the removed particulatesand sulfur oxides and/or sulfates are discharged through spent adsorberoutlet 86 and conveyed by gravity flow through spent adsorber line 118to the bottom of a spent adsorber regenerator comprising a lift piperiser 120 or transfer line. The spent adsorbers can be continuouslydischarged from the bottom of the granular bed filter and conveyed tothe regenerator lift pipe where they are regenerated, and/or cleansedbefore being recycled back to the granular bed filter. To this end, areducing gas (reduction gas) such as hydrogen, ammonia, carbon monoxide,or light hydrocarbon gases, such as methane, is injected upwardly intothe lift pipe riser by gas injector 122. The reducing gas can also bediluted with steam to attain a shift reaction or steam reforming inorder to produce hydrogen and carbon monoxide. The steam can be injectedinto the lift pipe riser (1) along and as part of the reducing gas or(2) through a separate steam injector 123. The reducing gas is injectedupwardly at a sufficient velocity and pressure to propel, carry,transport, and convey the adsorbers upwardly through the lift pipe riserinto an overhead collection vessel 124. In the lift pipe riser, thereducing gas reacts with the spent adsorbers and simultaneously removesthe particulates and sulfur oxides (SOx) while converting the sulfuroxides to hydrogen sulfide (H₂ S). Hydrogen produced by a steam shiftreaction or steam reforming serves as an effective and relativelyinexpensive reducing gas to convert sulfur oxides to hydrogen sulfides.The regenerated, cleansed adsorbers are recycled and conveyed from theoverhead vessel by gravity through regenerated adsorber line 114 intothe granular bed filter. Excess regenerated adsorbers can be removedfrom the system through overflow line 126 and discarded or stored in ahopper.

Methane can be an even more economical reductant or reducing gas undercertain conditions than hydrogen. When using methane in the lift piperiser, the reduction duration influences the sulfur dioxide (SO₂)pick-up capacity (regeneration) of the spent alumina adsorbers. At areduction temperature of 1,200° F. in the lift pipe riser, relativelyshort methane contact times of about 5 seconds are more effectivetowards restoring sulfur dioxide (SO₂) pick-up capacity (regeneration)of alumina adsorbers circulated in the granular bed filter and adsorberat a 1,300° F. adsorption temperature than longer methane contact timesof from 30 to 45 seconds. When the reduction temperature in the liftpipe riser is increased to at least 1,300° F., the effect of methanecontact time duration is negligible.

Desirably, the reducing gas also strips, scrubs, or otherwise removesthe captured nitrogen oxides (NOx) from the adsorbers in the lift piperiser and simultaneously converts the removed nitrogen oxides (NOx) towater or steam and molecular nitrogen (N₂) which can be safely vented tothe atmosphere from the overhead collection vessel through the outletgas line.

The effluent spent reducing gas, which contains hydrogen sulfide and theremoved particulates, is withdrawn from the overhead vessel 124 throughgas outlet line 128 where it is passed through one or more cyclones 130in order to remove most of the particulates via particulate dischargeline 132. The separated gases exit the cyclone through gas line 134where they can be fed to a bag house 136 to remove most of the remainingparticulates through particulate line 138. The filtered gases exit thebag house through gas line 140 where they are passed to an aminerecovery unit 142 to concentrate the hydrogen sulfide. Hydrogen sulfidefrom the vapor recovery and upgrading unit (not shown) downstream of thecatalytic cracker, can also be fed to the amine recovery unit. Theconcentrated hydrogen sulfide is passed from the amine recovery unitthrough concentrated hydrogen sulfide line 144 to a sulfur recovery unit146, such as a Claus plant, to recover elemental sulfur through sulfurrecovery line 148. The recovered sulfur can be safely stacked in pilesor transported elsewhere for other uses. If the level of hydrogensulfide (H₂ S) in the filtered spent reducing gas in line 140 issufficiently concentrated, the filtered gases can be sent directly tothe sulfur recovery unit 146 via bypass line 150, bypassing the aminerecovery unit.

In the preferred embodiment, in order to effectively and efficientlyremove the particulates and sulfur oxides (SOx) from theregenerator/combustion off-gases (flue gases), the off-gases shouldenter the granular bed filter at a temperature ranging from 200° F. to1,800° F. and most preferably from 500° F. to 1,400° F., at a pressurefrom atmospheric pressure to 500 psia. For best results, the granularbed filter should be operated at a temperature ranging from 200° F. to1,600° F., preferably from 1,000° F. to 1,400° F., and most preferablyfrom about 1,300° F. to about 1,350° F. at a pressure from 14 psia to300 psia and preferably from atmospheric pressure to 150 psia. Themaximum design operating temperature of the granular bed filter is2,000° F. The granular bed filter has an efficiency ranging from 85% to100% and preferably greater than 95%.

The solids flux flow rate of the adsorbers fed into the granular bedfilter is from 10 to 2,000 lbs/ft² hr, and preferably between 20 and 200lbs/ft² hr for best results. The regenerated and fresh adsorbers are fedinto the granular bed filter at a temperature ranging from 200° F. to1,800° F. and preferably from 500° F. to 1,400° F., at a pressureranging from 15 to 300 psia and preferably from atmospheric pressure to150 psia. The adsorbers range in diameter (size) from 1 mm to 13 mm andpreferably from 2 mm to 5 mm for best results. Adsorbers ranging in sizefrom 2 to 5 mm are not only effective in removing particulates butprovide excess capacity to adsorb sulfur oxides (SOx) and thereforeprovide a comfortable margin of safety to minimize downtime resultingfrom attrition or replacement of adsorbers. Only a small fraction ofalumina adsorbers, typically less than 1 % by weight, is utilized forsulfur dioxide (SO₂) capture. The low utilization of the aluminaadsorbers avoids the problem of alumina integrity. Integrity problemsarise when about 30% or more of the alumina adsorbers are used forsulfur dioxide (SO₂) capture in large amounts of steam.

The feed ratio (space velocity) of the sulfur oxide-removingcatalyst/promoter per lbs/min sulfur dioxide (SO₂) in the regeneratoroff-gases (flue gases) per lb of adsorber is from 1×10⁻³ to about 1×10⁻⁵and most preferably from about 2×10⁻⁴ to about 4×10⁻⁵ for best results.The ratio of catalyst/promoter to adsorbers by weight is in the range of1×10⁻⁶ :1 to about 1:3 and most preferably from about 2×10^(-6:1) toabout 1:9 for enhanced results.

The adsorbers can have a crush strength ranging from 1 to 10 lbs/mm andpreferably between 2 and 8 lbs/mm. The attrition weight of theregenerated adsorbers being recycled through the granular bed filter canrange from 0.1% to 2% and is preferably less than 1% per day for lessdowntime. The surface area-to-weight ratio of the adsorbers can rangefrom 5 to 400 m² /g unsteamed, and 2 to 250 m² /g if steamed duringpretreatment. The pore volume of the adsorbers can range from 0.3 to 1.5m² /g unsteamed, and preferably from 0.25 to 1 m² /g if steamed duringpretreatment. The pore radius of the adsorbers can range from 30 to 90 Åunsteamed and preferably from 50 to 200 Å if steamed duringpretreatment.

The bulk density of the moving bed of adsorbers can range from 20 to 120lbs/ft³ and preferably about 40 lbs/ft³. The bed of adsorbers movesdownwardly on the order of 1 to 30 in/hr and preferably from about 2 to20 in/hr. The flue gas residence time in the bed of adsorbers can rangefrom 1 to 10 seconds and preferably is about 2 seconds with asuperficial flue gas velocity through the bed ranging from 0.5 to 5ft/sec and preferably from about 1 to 2 ft/sec.

The solids residence time of the particulates as well as the adsorbersin the granular bed filter and scrubber is from 1 to 10 hours andpreferably from 2 to 4 hours for greater efficiency. The gas residencetime of the flue gases in the granular bed filter and adsorber is from 1to 5 seconds and preferably from 2 to 4 seconds for greatereffectiveness.

The lift pipe riser/adsorber-regenerator is preferably operated at atemperature of 1,000° F. to 1,600° F. and preferably from 1,200° F. to1,400° F., at a total pressure ranging from 15 to 300 psia andpreferably from atmospheric pressure to 150 psia, at a hydrogen partialpressure ranging from 0.1 p to 1 p and preferably from at least 0.5 pfor best results. The solids residence time of the particulates as wellas the adsorbers in the lift pipe riser can be from 15 seconds to 10minutes, preferably from 60 seconds to 150 seconds and the gas residencetime in the lift pipe riser can be from 10 to 30 seconds, preferablyfrom 16 to 18 seconds for best results. The spent adsorbers are heatedin the lift pipe riser to a temperature ranging from 800° F. to 1,600°F. and preferably from 1,200° F. to 1,400° F. for best results. The liftgas velocity in the lift pipe riser can range from 5 to 100 ft/sec andpreferably from about 20 to 40 ft/sec for best results.

The conversion level of removing particulates from the flue gas streamin the granular bed filter is from 85% to 100% and preferably at least95% for best results. The conversion level of removing sulfur oxides(SOx) from flue gases in the granular bed filter is from 85% to 100% andpreferably at least 95% for best results. The conversion level ofremoving nitrogen oxides (NOx) from flue gases in the granular bedfilter is from 85% to 100% and preferably at least 95% for best results.

The conversion level of removing particulates from the spent adsorbersin the lift pipe riser is from 90% to 100% and preferably from 95% to98% for better efficiency. The conversion level of converting sulfuroxides and/or sulfates to hydrogen (H₂ S) sulfide in the lift pipe riseris from 80% to 100% and preferably greater than 99% for greaterefficiency.

While the above operating conditions are preferred for best results, insome circustances it may be desirable to use other operating conditions.Furthermore, while the described granular bed filter is preferred tomost effectively remove particulates, sulfur oxides, and nitrogen oxidesfrom flue gases, in some circumstances it may be desirable to use othertypes of vessels, devices, or apparatus to simultaneously removeparticulates, sulfur oxides, and nitrogen oxides from flue gases, suchas those shown in U.S. Pat. Nos. 4,017,278; 4,126,435; and 4,421,038,which are hereby incorporated by reference in their entirety.

The sulfur oxide-capturing catalyst/promoter can be impregnated,deposited, or sprayed onto the adsorbers or fed separately with theadsorbers into the granular bed filter.

EXAMPLE 1

A sulfur dioxide (SO₂) adsorption capacity test was conducted with fluegas having an inlet composition of 1,000 ppmv sulfur dioxide, 3% byvolume molecular oxygen (O₂), and 2% by volume water vapor with a gasflow rate of about 10 cc/min at a temperature of 1,200° F. Aluminaadsorbers were used having a crush strength of 7.47 lbs/mm, an attritionrate of 0.06%, an unsteamed surface area of 198 m² /g, a pore volume of0.3609 cc/g unsteamed, and a pore radius of 32 Å unsteamed. The aluminaadsorbers removed 204 μl of sulfur dioxide (SO₂) per 50 mg of adsorbers.

EXAMPLE 2

A sulfur dioxide adsorption capacity test was conducted under the sameconditions as in Example 1, except that the alumina adsorbers had acrush strength of 1.75 lbs/mm, an attrition rate of 0.01%, a surfacearea of 269 m² /g unsteamed, a pore volume of 0.8426 cc/g unsteamed, anda pore radius of 38 Å unsteamed. The adsorbers removed 241 μl of sulfurdioxide (SO₂) per 50 mg of adsorbers.

EXAMPLE 3

A sulfur dioxide adsorption capacity test was conducted under theconditions of Example 1, except that the alumina adsorbers wereimpregnated with 2 ppm platinum catalyst/promoter to promote theadsorption of SOx. The platinum-promoted alumina adsorbed 270 μl sulfurdioxide (SO₂) per 50 mg of adsorbers.

EXAMPLE 4

A sulfur dioxide adsorption capacity test was conducted under theconditions of Example 2, except that the alumina adsorbers wereimpregnated with 2 ppm platinum. The platinum-promoted alumina adsorbersremoved 393 μl of sulfur dioxide (SO₂) per 50 mg of adsorbers.

EXAMPLE 5

A sulfur dioxide adsorption capacity test was conducted under theconditions of Example 1, except that the alumina adsorbers wereimpregnated with 6 ppm platinum. The platinum-promoted alumina adsorbersremoved 324 μl sulfur dioxide (SO₂) per 50 mg of adsorbers.

EXAMPLE 6

A sulfur dioxide adsorption capacity test was conducted under theconditions of Example 2, except that the alumina adsorbers wereimpregnated with 6 ppm platinum. The platinum-promoted alumina adsorbersremoved 414 μl of sulfur dioxide (SO₂) per 50 mg of adsorbers.

EXAMPLE 7

A regeneration test was conducted to regenerate the spentplatinum-promoted alumina adsorbers of Example 4, while simultaneouslyremoving the captured sulfur oxide (SOx) and/or sulfate from theadsorbers. The spent adsorbers were exposed to a pure dry hydrogenstream flowing at 10 cc/min for about 30 seconds at a temperature of1,200° F . The promoter was then subjected to an air purge to oxidizethe platinum sulfide to platinum. The regenerated adsorbers were thenused to adsorb the sulfur dioxide (SO₂) in the flue gas of Example 4 andachieved virtually a 100% sulfur dioxide (SO₂) removal rate in less than10 seconds.

EXAMPLE 8

A sulfur dioxide adsorption test was conducted on the flue gas ofExample 1 but at a temperature of 1,382° F. and using adsorberscomprising 100 mole percent magnesium (MgO) impregnated with 6% by wtceria (CeO₂). After 92 min., 26,300 μl of sulfur dioxide (SO₂) per 50 mgof adsorbers were adsorbed. The adsorbers had an activity of 0.748relative to the alumina adsorbers impregnated with ceria.

EXAMPLE 9

A sulfur dioxide adsorption test was conducted under the conditions ofExample 8, except that the ceria-impregnated adsorbers contained 92.6mole percent magnesium and 7.4 mole percent alumina. The adsorbersremoved 16,700 μl of sulfur dioxide (SO₂) per 50 mg of adsorbers and hada relative activity of 0.475.

EXAMPLE 10

A sulfur dioxide adsorption test was conducted under the conditions ofExample 8, except that the ceria-impregnated adsorbers contained 18.5mole percent magnesia and 8.5 mole percent alumina. The adsorbersremoved 10,550 μl of sulfur dioxide (SO₂) per 50 mg of adsorbers and hada relative activity of 0.3.

EXAMPLE 11

A sulfur dioxide adsorption test was conducted under the conditions ofExample 8, except that the ceria-impregnated adsorbers contained 55.8mole percent magnesia and 44.2 mole percent alumina. The adsorbersremoved 4,100 μl of sulfur dioxide (SO₂) per 50 mg of adsorbers and hada relative activity of 0.117.

EXAMPLE 12

A sulfur dioxide adsorption test was conducted under the conditions ofExample 8, except that the ceria-impregnated adsorbers contained 33.5mole percent magnesia and 66.5 mole percent alumina. The adsorbersremoved 1,700 μl of sulfur dioxide (SO₂) per 50 mg of adsorbers and hada relative activity of 0.048.

EXAMPLE 13

A sulfur dioxide adsorption test was conducted under the conditions ofExample 8, except that alumina adsorbers impregnated with 6% by weightceria were used. The adsorbers removed 650 μl of sulfur dioxide (SO₂)per 50 mg of adsorbers and had a relative activity of 0.018. The liquidhourly space velocity was 9,600 SCFH.

EXAMPLE 14

An attrition rate test was conducted with the adsorbers in Example 13.The adsorbers were found to have an attrition rate of 20.5%.

EXAMPLE 15

An attrition rate test was conducted with adsorbers comprising 16.7 molepercent magnesia and 83.3 mole percent alumina impregnated with 6 weightpercent ceria. The adsorbers were found to have an attrition rate of15.8%.

EXAMPLE 16

An attrition rate test was conducted with adsorbers comprising 15 molepercent magnesia and 50 mole percent alumina impregnated with 6 weightpercent ceria. The attrition rate was found to be 9.7%.

EXAMPLE 17

An attrition rate test was conducted with adsorbers containing 83.3 molepercent magnesia and 16.7 mole percent alumina. The attrition rate wasfound to be 7.0%.

In the system of FIG. 1, the sulfur oxides (SOx) and particulates are atleast partially removed by chemical adsorption, sometimes referred to asoxidative adsorption, with the captured SOx being converted to hydrogensulfide (H₂ S) when reacted with a reducing gas in the regenerator/liftpipe riser.

Captured sulfur dioxide (SO₂) reacts with alumina adsorbers to formalumina sulfate on the alumina adsorbers in accordance with thefollowing formula:

    3SO.sub.2 +1.5O.sub.2 +Al.sub.2 O.sub.3 →Al.sub.2 (SO.sub.4).sub.3 or Al.sub.2 O.sub.3.3SO.sub.3

The efficiency of chemical adsorption (oxidative adsorption) in removingsulfur dioxide (SO₂) from the flue gases in the granular bed filter isenhanced if the operating temperature of the granular bed filter is inthe range of 1,200° F. to 1,400° F.

The spent adsorbers containing the alumina sulfate are regenerated inthe lift pipe riser by reacting the spent adsorbers with a reducing gas,such as hydrogen, ammonia, carbon monoxide, or light hydrocarbon gases,such as methane, to remove and convert the alumina sulfate to hydrogensulfide (H₂ S). The regeneration of the adsorbers and the removal of thecaptured sulfur dioxide in the lift pipe riser is sometimes referred toas desorption.

The system of FIG. 2 is similar to the system of FIG. 1, except that theadsorbers remove the sulfur oxides (SOx) and particulates from the fluegas streams in the granular bed filter primarily by physical adsorption,sometimes referred to as non-oxidative adsorption, and the spentadsorbers are thermally regenerated in the lift pipe riser 120 by heat,preferably by combustion with an oxygen-containing combustion-sustaininggas, such as air, instead of a reducing gas, to remove the capturedsulfur oxides (SOx) and particulates from the spent adsorbers. Theadditional heat which is required for thermal regeneration can besupplied by combusting a fuel while in contact with the spent adsorbers.The fuel can be injected into the lower portion of the lift pipe riserthrough an auxiliary or supplemental fuel line 150. The fuel can betorch oil, hydrogen sulfide, or light hydrocarbon gases, such asmethane. Other fuels can be used. Air is injected upwardly into the liftpipe riser through air injector line 152 at a sufficient pressure andvelocity to convey, propel, carry, and transport the adsorbers alongwith the particulates and sulfur oxides to the overhead collectionvessel 124. During thermal regeneration, the particulates are removed(freed) from the adsorbers, the captured sulfur oxides are removed andbecome more concentrated, and/or the sulfates are removed from theadsorbers and converted to sulfur dioxide (SO₂). The effluent gasescontaining the concentrated levels of sulfur dioxide (SO₂), andsometimes hydrogen sulfide (H₂ S) from the combusted auxiliary fuel inthe overhead collection vessel are withdrawn from the vessel by gas line128 and fed to a sulfur recovery unit, such as a Claus plant to recoverelemental sulfur.

Adsorption efficiency, as measured by breakthrough times, for removingsulfur dioxide (SO₂) by physical adsorption (non-oxidative adsorption)is enhanced if the granular bed filter is operated at a temperatureranging from 500° F. to 800° F. The adsorption process for physicallyadsorbing sulfur dioxide on alumina adsorbers can be characterized bythe following formula:

    SO.sub.2 +Al.sub.2 O.sub.3 →Al(SO.sub.3).sub.3 or SO.sub.2.Al.sub.2 O.sub.3

The hydrogen sulfide (H₂ S) in the waste gases in the effluent line 128(FIG. 1) of the overhead collection vessel 124 can be removed andconcentrated by various methods, such as in an amine recovery unit 142with either diethanolamine (DEA) or monoethanolamine (MEA), the ironsponge process, or the hot potassium carbonate process. A DEA-operatingamine unit is preferred because it is more efficient and has lesschemical degradation and lower make-up risk than the other processes.

The amine recovery unit preferably decreases the concentration ofhydrogen sulfide in the waste gas stream to less than 1 part per cubicfoot of gas. DEA is preferred over MEA because of degradation of MEA bycarbonyl sulfide and carbon disulfide in the gases. DEA amine solutionswill absorb both hydrogen sulfide (H₂ S) and carbon dioxide (CO₂)according to the following reaction: ##STR1## Absorption of hydrogensulfide occurs in the amine recovery unit at 100° F. or below andrejection of sulfide is active at 240° F. The amine desulfurizationprocess which occurs in the amine recovery unit involves contacting thesour sulfur-containing gas stream (waste gases) with a cool DEA aminesolution to absorb the hydrogen sulfide and then regenerate the amineand strip the hydrogen sulfide from the amine solution by heating.

In the preferred embodiment, the amine recovery unit takes the formshown in FIG. 3, although other types of amine recovery units can beused, if desired. In the embodiment of FIG. 3, sour waste gases (acidgases) in waste gas line 140 are fed to an inlet scrubber 200 whichremoves (scrubs) entrained liquids, including distillate and water, fromthe waste gases. The scrubbed gases are discharged from the inletscrubber through scrubber discharge line 202 and fed to the bottomportion of a contactor or absorber column 204. A DEA amine feed ispumped into the top portion of the absorber column by amine charge pump206 via amine feed line 208. In the adsorber column, the scrubbed wastegases are contacted in countercurrent flow relationship with the aminefeed to react the hydrogen sulfide and the carbon dioxide in the wastegases with the amine. The adsorber column can be a trayed or packedtower and provides gas-liquid adsorption.

Rich amine is discharged from the bottom of the adsorber column throughrich amine line 210 and fed to a flash tank 212 where it is flashed at areduced pressure to remove entrained gases through entrained gas line214. The flashed rich amine gases are discharged from the bottom of theflash tank through discharge line 216 and fed to and filtered in acarbon filter 218. The filtered rich amine is fed through filtered amineline 220 to a rich/lean amine heat exchanger 222 where the rich amine isheated. The heated rich amine is discharged from the heat exchangerthrough heated rich amine line 224 and fed to the upper portion of astripper column, steam stripper, or still 226.

Steam is injected into the lower portion of the steam stripper 226through steam injection line 228. In the steam stripper, the rich aminesolution is regenerated and stripped of acid gases by the steam. Theconcentrated acid gases are withdrawn from the steam stripper throughoverhead acid gas line 230 and cooled in a water cooler or condenser232. The cooled acid gases are passed through cooled acid gas line 234and collected in a reflux accumulator 236. Part of the concentrated acidgases in the reflux accumulator can be recycled, refluxed, and pumpedinto the upper portion of the steam stripper (stripper column) 226 byreflux pump 238 via reflux lines 240 and 242. The excess acid gases canbe discharged from the reflux accumulator through excess gas line 244and flared or sent to a sulfur recovery unit, such as a Claus plant.

The stripped rich amine is discharged from the bottom of the steamstripper 226 through stripped amine line 246 and fed to a reboiler 248.Steam is boiled out of the amine in the reboiler and withdrawn throughoverhead steam line 228 where it is injected into the lower portion ofthe steam stripper 226. The residual boiled lean amine is dischargedfrom the reboiler through lean amine discharge line 250 and passedthrough lean amine line 252 to heat exchanger 222. The inventory of thelean amine in amine line 252 is controlled by surge tank 254.

Lean amine exits the heat exchanger 222 through outlet line 256 and ispumped through line 258 to a cooler or heat exchanger 260 by boosterpump 262. The lean amine solution is cooled in the heat exchanger 260.The cooled amine is discharged from the heat exchanger 260 throughcooled amine line 264 and pumped through lean amine feed line 208 intothe upper portion of the absorber column 204 by amine charge pump 206.

Effluent gases are withdrawn from the absorber column through overheadgas line 266 and fed to an outlet scrubber 268. The outlet scrubberscrubs the gases from the gas line 266 to recover any residual aminesolution carried over in the effluent gases. The sweet scrubbed gasesare discharged from the outlet scrubber through sweet gas line 270.

The acid waste gases in lines 140 and 150 (FIG. 1), as well as the acidgases in acid gas line 244 from the reflux accumulator 236, are fed to asulfur recovery unit and scavenger plant, preferably a Claus plant, suchas the type shown in FIG. 4. The Claus plant can recover 99.0% or moreof the elemental sulfur in the influent acid gases.

As shown in FIG. 4, acid gases enter an oxidation unit and waste-heatboiler 300 through an acid gas inlet line 302. In the oxidation unit,about one-third of the hydrogen sulfide (H₂ S) in the acid gases isoxidized to sulfur dioxide (SO₂) and water or steam in accordance withthe following exothermic reaction:

    H.sub.2 S+3/2 O.sub.2 →SO.sub.2 +H.sub.2 O

The reaction furnace section 306 of the unit 300 is downstream of theburner and provides a thermal region in which about 70% by weight of thehydrogen sulfide (H₂ S) of the remaining acid gases and the sulfurdioxide (SO₂) is converted to elemental molecular sulfur and water orsteam in accordance with the following endothermic reaction:

    2H.sub.2 S+SO.sub.2 →3/2 S.sub.2 +2H.sub.2 O

Water is fed into the boiler section 308 of the unit 300 through waterline 310. The hot reaction gases in the reaction furnace, which can beat a temperature such as 2,300° F., are cooled by the water in the waterpipes of the boiler section to a much cooler temperature, such as 1,100°F. The water in the water pipes of the boiler section is boiled andheated by the hot reaction gases and converted to steam. Steam isremoved from the boiler section through steam line 312. In the boilersection, the elemental sulfur is converted to S₆ and S₈ in accordancewith the following exothermic reactions:

    S.sub.2 →1/3 S.sub.6 and S.sub.2 →1/4 S.sub.8

Hot gases containing S₆ and S₈ are withdrawn from the unit through gasbypass line 314.

The partially stripped reaction gases are removed from the unit 300through outlet gas line 316. The stripped reaction gases typicallycontain hydrogen sulfide (H₂ S), sulfur dioxide (SO₂), elemental sulfur(S₂), nitrogen (N₂), carbonyl sulfide (COS), carbon disulfide (CS₂), andsteam. The stripped reaction gases can be withdrawn from the unitthrough gas line 316 at a temperature ranging from 550° F. to 600° F.The stripped reaction gases in the gas line 316 are fed to a heatexchanger 318 to cool the reaction gases to about 530° F. and condenseor precipitate some of the sulfur. The condensed or precipitated sulfuris removed from the heat exchanger through sulfur line 320.

The cooled reaction gases are withdrawn from the heat exchanger 318through cooled gas line 322 and fed to a first Claus converter 324. Thebottom portion of the converter contains a fixed catalyst bed 326 ofsulfur-capturing catalysts. The reaction gases are passed through thefixed catalyst bed in the first converter to catalytically react theremaining hydrogen sulfide (H₂ S) with the sulfur dioxide (SO₂) to formwater and free sulfur. The products are heated by the catalytic reactionto over 650° F. The reaction products are discharged from the firstconverter through discharge line 328 and cooled in a cooler or heatexchanger 330 to condense, precipitate, and/or recover more sulfur. Thesulfur is removed from the heat exchanger through sulfur recovery line332.

The cooled reaction gases, which can be cooled to below 400° F., arewithdrawn from the heat exchanger 330 through cooled reaction gas line334 and fed to a second Claus converter 336. The second Claus converteralso has a fixed catalyst bed 338 of elemental sulfur-capturingcatalyst. The reaction products are passed through the catalyst bed 338to catalytically react the remaining hydrogen sulfide (H₂ S) with thesulfur dioxide (SO₂) to form water and free sulfur. The resultingreaction products are heated to a temperature slightly below 500° F. bythe reaction in the second converter. The reaction products aredischarged from the second converter through reaction product outletline 339 and cooled in a cooler or heat exchanger 340 to condense,precipitate, and/or remove substantially all of the remaining sulfur.

The sulfur is removed from the heat exchanger 340 through sulfurrecovery line 342. The cooled tail gases are withdrawn from the heatexchanger 340 through tail gas outlet line 344 and passed to tail gasclean-up equipment 346, such as Beavon and Stretford processingequipment, to clean up the tail gases. The sweet cleansed tail gases arewithdrawn from the tail gas clean-up equipment through sweet gas line348. Sulfur recovered from the tail gases are removed from the tail gasclean-up equipment through sulfur recovery line 350.

While the above two-stage Claus plant is preferred because it recoversat least 95% elemental sulfur, other types of Claus plants can be used,if desired, such as a split-stream Claus plant, a partial-oxidationClaus plant, an ultra three-stage Claus plant, etc.

Although embodiments of this invention have been shown and described, itis to be understood that various modifications and substitutions, aswell as rearrangements and combinations of parts, components, and/orequipment can be made by those skilled in the art without departing fromthe novel spirit and scope of this invention.

What is claimed is:
 1. A gas purification system, comprising:a vesselcomprising a housing defining and positioned about a vertical axis, aplurality of conduits connected to and communicating with said housing,said conduits including a gas inlet, a gas outlet, an adsorber inlet andan adsorber outlet, said gas inlet comprising a downwardly facing gasinlet being positioned at an angle of inclination ranging from about 30°to about 90° relative to a horizontal axis for feeding and directing theflow of influent gases generally downwardly through said housing, anannular frustoconical adsorber collection reservoir positioned withinsaid housing, said reservoir having flared sidewalls extending below theheight of said gas inlet, and discharge chutes communicating with saidreservoir comprising substantially symmetrical inclined chutes extendingdownwardly and outwardly from said reservoir at an angle of inclinationrelative to said vertical axis in a direction away from said gas inlet;sulfur oxide and particulate removal balls located in said reservoir,inclined chutes, and housing for simultaneously removing a substantialportion of particulates and sulfur oxides from a gaseous stream in saidhousing, said sulfur oxide and particulate removal balls entering andexiting said housing through said absorber inlet and absorber outlet,respectively; and said sulfur oxide and particulate removal ballsincluding a bed of sulfur oxide-capturing and particulate-removing ballsselected from the group consisting essentially of adsorbers andadsorbers with at least one promoter thereon, said adsorbers beingsubstantially present in the absence of a carbonaceous aggregate andcomprising an oxide of at least one metal selected from the groupconsisting of aluminum, bismuth, manganese, yttrium, antimony, tin, arare earth metal, a Group 1a metal, and a Group 2a metal, and saidpromoter selected from the group consisting essentially of a rare earthmetal, a Group 8 metal, chromium, vanadium, rhenium, and combinationsthereof.
 2. A catalytic cracking and gas purification system,comprising:a catalytic cracker for cracking a hydrocarbon feedstock inthe presence of a catalytic cracking catalyst; a cracking catalystregenerator positioned downstream of and communicating with saidcatalytic cracker for regenerating spent catalytic cracking catalyst andemitting regenerator off-gases containing sulfur oxides andparticulates; a granular bed filter comprising an elongated,substantially upright, sulfur oxide-capturing and particulate-removingvessel, said vessel positioned downstream and communicating with saidcracking catalyst regenerator for substantially filtering saidparticulates and substantially removing said sulfur oxides from saidregenerator off-gases to form substantially purified gases having asubstantially lower concentration of sulfur oxides and particulates thansaid combustion off-gases, said vessel having a roof, a frustoconicalbottom portion having flared walls converging generally downwardly, andsidewalls connecting said roof and said bottom portion, said vesseldefining a vertical axis and containing a downwardly moving bed of ballscomprising sulfur oxide-capturing and particulate-removing material,said vessel having an elongated upright conduit defining a substantiallyvertical gas inlet extending downwardly through said roof along saidvertical axis into said bed for passing said regenerator off-gases intosaid bed, said vessel having a purified gas outlet line having an intakemouth located above said bed and a frustoconical collection reservoirwith flared sidewalls extending below the mouth of said purified gasoutlet line, said flared sidewalls diverging generally downwardly fromand cooperating with said roof for holding balls comprising regeneratedsulfur oxide-capturing and particulate-removing material, said reservoirpositioned above said bed and circumferentially surrounding a portion ofsaid upright conduit, and a plurality of chutes extending between andcommunicating with said reservoir and said bed for feeding said ballscomprising said regenerated sulfur oxide-capturing andparticulate-removing material from said reservoir to said bed, saidchutes including a vertical chute concentrically position about aportion of said conduit below said reservoir and a plurality ofsubstantially symmetrical inclined chutes positioned at an angle ofinclination relative to the vertical axis of said vessel in a directionaway from said vertical gas inlet; said sulfur oxide-capturing andparticulate-removing material selected from the group consistingessentially of adsorbers and adsorbers with at least one promoterthereon, said adsorbers substantially comprising an oxide of at leastone metal selected from the group consisting of aluminum, bismuth,manganese, yttrium, antimony, tin, a rare earth metal, a Group 1a metal,and a Group 2a metal, and said promoter selected from the groupconsisting essentially of a rare earth metal, a Group 8 metal, chromium,vanadium, rhenium, and combinations thereof; an upright elongated, liftpipe riser spaced laterally from and communicating with said granularbed filter for regenerating spent balls comprising spent sulfuroxide-capturing and particulate-removing material, said lift pipe riserspaced laterally away from said cracking catalyst regenerator, said liftpipe riser having a lower portion and an upper portion, an overheadcollection vessel connected to and communicating with said upperportion, and injector means for injecting a gas selected from the groupconsisting essentially of a reducing gas and an oxygen-containing,combustion-sustaining gas, into said lower portion of said lift piperiser with a sufficient velocity and pressure to propel and carry saidspent material generally upwardly through said lift pipe riser into saidoverhead collection vessel, said reducing gas selected from the groupconsisting essentially of hydrogen, ammonium, carbon monoxide, lighthydrocarbon gases, and combinations thereof; a spent material feedlineextending between and connecting said frustoconical bottom portion ofsaid sulfur oxide-capturing and particulate-removing vessel to thebottom portion of said lift pipe riser for feeding said spent ballscomprising spent sulfur oxide-capturing and particulate-removingmaterial from said sulfur oxide-capturing and particulate-removingvessel to said life pipe riser; a regenerated material feed lineextending between and connecting said overhead collection vessel to saidroof of said sulfur oxide-capturing and particulate-removing vessel forfeeding regenerated balls comprising said regenerated material from saidoverhead collection vessel to said reservoir of said granular bedfilter; solid-gas separation means positioned downstream andcommunicating with said overhead collection vessel of said lift piperiser for substantially removing particulates from effluent gasesemitted from said overhead collection vessel; and sulfur recovery meanspositioned downstream and communicating with said solid-gas separationmeans for recovering elemental sulfur from said effluent gases.
 3. Acatalytic cracking and gas purification system in accordance with claim2 wherein said catalytic cracker comprises a fluid catalytic cracker. 4.A catalytic cracking and gas purification system in accordance withclaim 2 wherein said regenerator includes means for substantiallycompletely burning carbon monoxide in said regenerator.
 5. A catalyticcracking and gas purification system in accordance with claim 2 whereinsaid solid-gas separation means include at least one cyclone.
 6. Acatalytic cracking and gas purification system in accordance with claim2 wherein said solid-gas separation means includes a bag house.
 7. Acatalytic cracking and gas purification system in accordance with claim2 wherein said adsorbers are selected from the group consistingessentially of gamma alumina, chi-eta-rho alumina, delta alumina, andtheta alumina.
 8. A catalytic cracking and gas purification system inaccordance with claim 7 wherein said adsorbers comprise said adsorberswith at least one promoter and said at least one promoter comprises 2ppm to 6 ppm by weight of said adsorbers and is selected from the groupconsisting essentially of platinum, ceria, and combinations thereof. 9.A catalytic cracking and gas purification system in accordance withclaim 2 wherein said adsorbers comprise magnesium aluminate spinels. 10.A catalytic cracking and gas purification system in accordance withclaim 2 wherein said adsorbers comprise alumina and magnesia.
 11. Acatalytic cracking and gas purification system in accordance with claim10 wherein said adsorbers comprise said adsorbers with at least onepromoter and said at least one promoter comprises 2 ppm to 6 ppm byweight of said adsorbers and is selected from the group consistingessentially of platinum and ceria.
 12. A catalytic cracking and gaspurification system in accordance with claim 2 wherein said injectormeans comprises a regenerating reducing gas injector for injecting aregenerating reducing gas into said lift pipe riser and a steam injectorfor injecting steam into said lift pipe riser.